Catalysts for producing organic carbonates

ABSTRACT

A process for producing various organic carbonates by performing transesterification and disproportionation reactions in dual vapor/liquid phase mode preferably in the presence of solid catalyst composition selected from the group consisting of oxides, hydroxides, oxyhydroxides or alkoxides of two to four elements from Group IV, V and VI of the Periodic Table supported on porous material which has surface hydroxyl groups and the method of reactivating catalyst deactivated by polymer deposition by contacting the deactivated catalyst with a solution of hydroxy containing compound in a solvent such as benzene or THF.

CROSS-REFERENCE TO RELATED APPLICATION

This application is a continuation of U.S. patent application Ser. No.11/957,256, filed Dec. 14, 2007, which is a divisional application ofU.S. patent application Ser. No. 11/256,394, filed Oct. 21, 2005, nowU.S. Pat. No. 7,378,540, the disclosure of which is hereby incorporatedby reference.

BACKGROUND OF DISCLOSURE

1. Field of the Disclosure

The present invention relates to the process of producing variousorganic carbonates by performing chemical reactions, which are limitedby equilibrium, and separating various chemical compounds involved.Achieving higher conversion than equilibrium condition is highlydesirable for better economic reward. This invention discloses a methodof shifting the equilibrium position of a chemical reaction to achievehigher conversion.

2. Background

Diaryl carbonates, for example diphenyl carbonate (DPC), are animportant raw material for the production of polycarbonates. The currentstate of the art produces DPC from dimethyl carbonate (DMC, (CH₃O)₂C═O)and phenol in two steps by employing multiple reactors and a homogeneouscatalyst, such as titanium alkoxide. The following three reactions areinvolved in producing DPC ((C₆H₅O)₂C═O).

The equilibrium for each of the above reactions lies on far left side,whereas it is an object of the present invention that the equilibriumsare moved to the right.

J. L R. Williams et al. (J. Org. Chem., 24 (1) pp. 64-68, 1959)discovered that disproportionation of unsymmetrical carbonates such asmethyl phenyl carbonate (MPC) and symmetrical carbonates such asdibenzyl carbonate can be performed in the presence of a suitablehomogeneous catalyst, particularly metal alkoxide. However, the reactionaccompanies a number of undesired side reactions such as decompositionof carbonate compounds to carbon dioxide, polymerization, formation ofolefins, and ethers. The authors conclude that the course of thereactions depends on the structure of the carbonates and on the catalystand the more alkaline catalysts promote more side reactions for thedisproportionation of alkyl phenyl carbonates than mildly acidiccatalyst such as titanium butoxide. Alkyl phenyl ether is the majorundesired by-product.

U.S. Pat. No. 4,045,464 (1977) discloses a process for preparing diarylcarbonates from phenyl alkyl carbonates via the Reaction (3) in thepresence of homogeneous Lewis acid catalysts of the formula AIX₃, TiX₃,UX₄, TiX₄, VOX₃, VX₅, and SnX₄, where X is a halogen, acetoxy, alkoxy oraryloxy group. DPC was produced by performing the disproportionation ofethyl phenyl (EPC) or MPC at 180° C. with 95% selectivity in thepresence of homogeneous titanium catalysts.

U.S. Pat. No. 4,554,110 (1985) discloses an improved process for thepreparation of aromatic carbonates from a dialkyl carbonate and phenolin the presence of catalyst comprising polymeric tin compounds. Diarylcarbonates are prepared by performing disproportionation of alky arylcarbonate.

U.S. Pat. No. 5,210,268 (1993) discloses a process for producing diarylcarbonates by performing various transesterification reactions in tworeaction zones. An aromatic carbonate mixture is prepared by atransesterification reaction between a dialkyl carbonate, an alkyl arylcarbonate and a mixture thereof, and an aromatic hydroxy compound in thefirst step. A diaryl carbonate is produced by primarily performingdisproportionation of alkyl aryl carbonate in the second step. Thepatent discloses the method of overcoming unfavorable equilibrium. Thetransesterification is performed in reactive distillation mode byfeeding reactants to continuous multistage distillation columns in thepresence of a homogeneous catalyst or a solid catalyst, whilecontinuously withdrawing the produced aromatic carbonate mixture as ahigh boiling point product from a lower portion of the distillationcolumn and continuously withdrawing light coproducts such as aliphaticalcohol or dialkyl carbonate as vapor stream from upper portion of thedistillation columns. The disproportionation of alkyl aryl carbonate todiaryl carbonate and dialkyl carbonate is also performed in similarfashion. Diaryl carbonate is continuously produced by performingtransesterification and disproportionation in sequence, utilizingmultiple multi-stage distillation columns.

U.S. Pat. No. 5,872,275 (1999) and U.S. Pat. No. 6,262,210 (2001)disclose for the process producing diaryl carbonate from dialkylcarbonate and an aromatic hydroxy compound in the presence of liquidhomogeneous catalyst and methods removing heavy high boiling pointbyproducts and regenerating catalyst for recycle.

U.S. Pat. No. 6,093,842 (2000) discloses the process producing diarylcarbonate from a dialkyl carbonate, an aromatic hydroxy compound and amixed solution containing alkyl aryl carbonate by introducing threereactant streams into an extractive/reactive distillation column in thepresence of a catalyst. Examples of the catalysts are lead compounds,copper compounds, alkali metal compounds, nickel compounds, zirconiumcompounds, titanium compounds, vanadium compounds, etc. The byproductsare CO2, anisole, benzoates and heavy materials.

U.S. Pat. No. 5,426,207 (1995) discloses the process producing a diarylcarbonate such as DPC by conducting transesterification of DMC with anaromatic hydroxy compound and disproportionation of alkyl aryl carbonatein presence of a homogeneous catalyst in three successive reactionzones. Conditions are selected to maximize formation of alkyl arylcarbonate in the first and second reaction zones, whiledisproportionation is favored in the third reaction zone.

U.S. Pat. No. 6,767,517 (2004) discloses a process for the continuousproduction of diaryl carbonates. The process uses three reactivedistillation column reactors and two rectification columns for theseparation of intermediate reaction product and final product.

To alleviate shortcomings associated with using homogeneous catalystsfor the production of diaryl carbonates, U.S. Pat. No. 5,354,923 (1994)and U.S. Pat. No. 5,565,605 (1996), and WO 03/066569 (2003) discloseheterogeneous catalysts.

WO 03/066569 discloses a process continuously producing an aromaticcarbonate such as DPC in the presence of a heterogeneous catalystprepared by supporting titanium oxide on silica (3 mm diameter) in twosteps.

JP 54-125,617 (1979) and JP Application No. 07-6682 [1995] discloseheterogeneous catalysts for the preparation of diphenyl carbonate bytransesterification of DMC with phenol to MPC and disproportionation ofMPC to DPC in the presence of MoO₃ or V₂O₅ supported on an inorganicsupport such as silica, zirconia or titania. The transesterification anddisproportionation are carried out in a reactor-distillation towerconsisting of a reactor and distillation tower with removal of theco-products by distillation.

The publication by Z.-H. Fu et al. (J Mol. Catal. A: ChemicaM 18, (1997)pp. 293-299) reports the synthesis of diphenyl carbonate from DMC andphenol in the presence of various heterogeneous metal oxide catalysts.The best selectivity is reported for the MoO₃ (20 wt. % optimum loading)catalyst supported on silica.

Due to many shortcomings of current DPC processes, an improved processis highly desired for saving materials, cheaper construction cost,consuming less energy, and plant operation cost. Although the prior artdoes not address catalyst life, all heterogeneous catalysts eventuallydeactivate and become useless. Heterogeneous catalysts are onlypractical in commercial use, if their cycle and service times are longenough or the catalysts can be rejuvenated in situ without seriousfinancial cost. Thus, the issues of the catalyst deactivation and methodof regeneration remain as substantial bars to the commercial applicationof the prior art.

SUMMARY OF THE DISCLOSURE

This invention relates to a process for producing various organiccarbonates by performing transesterification and disproportionationreactions in the presence of solid catalyst in dual phase reactor. Thepresent process is beneficial for performing equilibrium limitedchemical reactions involving organic carbonates, where both gas phaseand liquid phase must coexist to shift the equilibrium position to rightside of an equilibrium reaction, thereby resulting in high conversion.The chemical reactions are performed in the presence of one or multipleheterogeneous catalysts.

The preferred solid catalysts disclosed in this invention are mixedoxide catalysts composed of two to four different elements from GroupIV, V and VI of the Periodic Table, preferably Ti, Zr, Hf₁ Nb, Ta, Mo,V, Bi and Si supported on porous materials such as silica, which havesurface hydroxyl groups. Supported metal alkoxide or mixed metalalkoxide catalysts of the Group IV and V metal alkoxides, such astitanium alkoxides, zirconium alkoxides, vanadium alkoxides, niobiumalkoxides, VO(OR)₃ or oligomers of oxoalkoxide, and the like constitutea preferred catalyst group. The transesterification catalyst and thedisproportionation catalyst utilized in the present process may be thesame or different.

In general, heterogeneous catalysts are more desirable compared tohomogeneous catalyst, because of difficulties involved in recyclinghomogeneous catalysts. However, all heterogeneous catalysts eventuallydeactivate. Deactivated catalysts must be replaced either with freshcatalyst or regenerated insitu without too much difficulty.

For the purposes of the present invention, the term “dual phase mode”means any process having both a liquid and a vapor phase present in thereaction zone regardless of the means of achieving the vapor and liquidphases including “reactive distillation”, “catalytic distillation”,boiling and concurrent reaction and fractional distillation in a column.The term “treated support” or “treated silica” is understood to mean asupport having an optimized population of surface hydroxyl groups for agiven surface area for the preparation of the catalysts describedherein.

In a preferred embodiment the process for the production of diarylcarbonate comprises:

-   -   (a) a plurality of reaction zones comprising a primary and a        secondary reaction zone;    -   (b) supplying to said primary reaction zone a dialkyl carbonate        and an aromatic hydroxy compound;    -   (c) maintaining the primary reaction zone under dual phase and        reaction conditions conducive for the formation of alkyl aryl        carbonate;    -   (d) transesterifying the dialkyl carbonate with the aromatic        hydroxy compound in the presence of a solid catalyst selected        from the group consisting of two to four elements from Group IV,        V and VI of the Periodic Table supported on porous material        which have surface hydroxyl groups;    -   (e) recovering a dual phase product stream from the primary        reaction zone;    -   (f) separating the dual phase product stream from (e) to recover        vaporous alkyl alcohol and liquid alkyl aryl carbonate;    -   (g) maintaining the secondary reaction zone under dual phase and        reaction conditions conducive for the disproportionation of        alkyl aryl carbonate to diaryl carbonate;    -   (h) disproportionating the alkyl aryl carbonate in the presence        of a solid catalyst selected from the group consisting of two to        four elements from Group IV, V and VI of the Periodic Table        supported on porous material which have surface hydroxyl groups;    -   (i) recovering a dual phase product stream from the secondary        reaction zone; and    -   (j) separating the dual phase product stream from (i) to recover        a vaporous component comprising aryl alkyl carbonate and liquid        product comprising diaryl carbonate.

More preferably additional steps are carried out:

-   -   (k) separating diaryl carbonate by distillation; and    -   (l) recycling the aryl alkyl carbonate to the secondary reaction        zone.

It is important that a support should have surface hydroxyl groups.Silica is a preferred support. Preferably the support is a treatedsupport.

Other aspects and advantages will be apparent from the followingdescription and the appended claims.

BRIEF DESCRIPTION OF DRAWINGS

The FIGURE is schematic representation of one embodiment of the processof the present invention.

DETAILED DESCRIPTION

The process improvements were made by introducing dual phase reactors toperform transesterification and disproportionation to establish thefollowing process conditions for the production of DPC:

-   -   (1) feeding essentially an ethanol-free stream to the        transesterification reactor;    -   (2) minimizing DEC in the feed stream to the disproportionation        reactor;    -   (3) performing the transesterification in the presence of phenol        such that the phenol to EPC mole ratio be greater than 0.2,        preferably greater than about 0.3, most preferably greater than        about 0.35;    -   (4) performing the disproportionation in the presence of phenol        at phenol to DPC mole ratio of in the range of 0.05 to 10,        preferably from 0.1 to 6 and phenol to EPC mole ratio of from        0.01 to 6;    -   (5) optionally introducing a trace amount of water into the        disproportionation catalytic reaction zone in an amount up to        0.3 wt %, preferably up to 0.1 wt %,    -   (6) capitalizing on the absence of a DEC/ethanol azeotrope to        separate DEC from ethanol;    -   (7) essentially eliminating cold spots in the catalytic reaction        zones;    -   (8) eliminating catalyst recycle, catalyst separation, and        continuous addition of make-up catalyst;    -   (9) effective mixing of reactants and products; and    -   (10) preferably maintaining the presence of both vapor and        liquid phases in the reaction zones to remove a low boiling        reaction product into a gas phase to obtain high conversion, not        limited by equilibrium constant.

The fixed bed reactors are operated in the dual phase mode, which meanscoexistence of both vapor and liquid in the reaction zones creating aboiling condition in the catalytic reaction zone of a dual phase reactoris desirable. Nevertheless, operating dual phase reactor under boilingcondition is not necessary as long as both gas and liquid phases coexistin the reaction zone.

The flow direction in a dual phase reactor can be either down-flow orup-flow. Down-flow is preferred. Also the fixed bed dual phase reactorshave recycle loops. A down-flow boiling point reactor is operatedpreferably to have a negative pressure drop, which means a lowerpressure at the bottom of the catalyst bed than the top of the catalystbed. The negative pressure drop is created by high mass flow rate ofvapor-liquid mixture through a fixed heterogeneous catalyst bed incatalytic reaction zone. The negative pressure drop is desirable but notnecessary. A pressure drop of approximately 0.2 psi per ft or more ispreferable.

Operation of a dual phase reactor under boiling point condition hasseveral advantages over the operation of a traditional fixed bedreactor, namely:

-   -   mass transport of the reactants from the bulk phase into the        catalyst particles and transport of the reaction products from        the inside pores of shaped catalyst pallets in the catalytic        reaction zone;    -   lateral mixing of the reactants and products in bulk phase;    -   development of no or negligible temperature drop or hot spot in        a catalytic reaction zone for either endo- or exothermic        reaction, and    -   removing light reaction products from the liquid reaction medium        into gas phase, which results in a favorable equilibrium        direction.

At a given reaction temperature and flow rate in the catalytic reactionzone, the boiling reaction condition in catalytic reaction zone iscreated by controlling the pressure, which is determined by thecomposition of the reaction medium. Optionally one may choose to use alower boiling solvent to create better boiling condition in thecatalytic reaction zone, when the boiling point of reactant or productis close to or higher than intended reaction temperature.

This invention is particularly useful for reactions where the boilingpoint of reaction product is higher than reaction temperature or tooclose to reaction temperature such that there will be no or insufficientgas phase volume in the reaction zone under the reaction condition.Examples of such a reaction include producing diphenyl carbonate bydisproportionation of ethylphenyl carbonate or methyl phenyl carbonateor producing bis-(2-ethyl-1-hexyl)carbonate by transesterificationreaction of an alkyl carbonate with 2-ethyl-1-hexanol. The preferredreaction temperature is from about 200° to 400° F. Since DMC and DECboil at temperatures higher than about 305° F., creating a vapor phasein large commercial fixed bed reactor becomes difficult at relativelyhigh liquid flow rates. By controlling the flow of vapor volume to thereactor at a given flow rate of liquid reaction mixture in a dual phasereactor as disclosed in this invention, it becomes relatively easy tocreate the proper volume of gas phase in the reaction zone at givenreaction conditions and maintaining a steady state reactor operation.

Significant advantages obtained over prior art are:

(a) high productivity of desired products,

(b) high selectivity of desired products,

(c) no separation of catalyst from the reaction product stream,

(d) excellent catalyst life,

(e) less energy consumption, and

(f) easy regeneration of deactivated catalyst.

The feed to transesterification reactor preferably comprises freshphenol and DEC, and a mixed recycle stream containing DEC, PhOH and EPC.An essentially ethanol-free feed stream to transesterification reactorcreates favorable conditions for the conversion of phenol and a highreaction rate. This is also true for essentially DEC-free feed stream todisproportionation reactor. There are practically no cold spots in thecatalytic reaction zones, which is important for maintaining steadyreaction rates. Since there are both vapor and liquid phases intransesterification and disproportionation reaction zones in dual phasereactors, light reaction products (ethanol and DEC in each reactionzone) are vaporized to a gas phase which creates favorable conditionsfor high productivity of the intended products. As there is no exposureof reactants or products to unnecessarily high temperature in thepresence of catalyst, there are less undesired by-products, making iteasier and cheaper to purify crude DPC product stream. Because the feedstream to disproportionation reaction zone may be composed of fairlyhigh boiling compounds, such as EPC or MPC, nitrogen, a low boilingcomponent or both are introduced into the reaction zone to createsufficient volume of gas phase in the reaction zone. Examples such lowboiling components are ethyl ether, propyl ether, dimethyl ether,methane, ethane, propane, butane, hexane, heptanes, toluene, andxylenes.

Preferably all or at least substantial portion of low boiling componentsare introduced into the reaction zone as super heated gas. Since ethanoland DEC do not form an azeotrope, DEC separation for recycle results inlower energy consumption and lower construction cost compared withexisting commercial processes where DPC is produced from DMC and phenol.

There are two or three reaction zones. In the first reaction zone, analkyl aryl carbonate such as EPC is produced by transesterification ofDEC with phenol. In the second reaction zone, DPC is produced byperforming disproportionation of EPC. The reaction zones comprise twofixed bed reactors. Optionally three reaction zones may comprise threefixed bed reactors or two fixed bed reactors and a catalyticdistillation column reactor. The reactors are loaded with one or twodifferent heterogeneous catalysts.

Examples of organic carbonates produced by this invention are DPC, EPC,MPC, DEC, DMC, bis-(2-ethylhexyl) carbonate; and fatty acidmono-glyceride carbonates. The present invention is particularly usefulin producing diphenyl carbonate (DPC) from DEC and phenol. Although DECis the preferred dialkyl carbonate for the production of DPC, it isunderstood that the process is also useful for the production of DPC byusing DMC or any other alkyl carbonate or alkyl aryl carbonate.

To avoid excessive heating of the reboiler while conductingtransesterification and disproportionation in catalytic distillationcolumn reactor, a low boiling component or a mixture of low boilingcomponents may be pumped directly into the reboiler. Therefore, thistechnique is a part of the present invention for production of variousorganic carbonates for both solid catalysts and homogeneous catalystused in a process. This technique has not been disclosed in the priorart for the production of organic carbonates such as DPC, DEM, and DEC.

The preferred heterogeneous catalysts are supported mixed oxides,hydroxides, oxyhydroxides and alkoxides of Group IV, V and VI elementswhich are deposited on porous supports. The mixed oxide catalysts may becombinations of two, three or four elements chosen from Mo, Nb, Ti, VZr, Bi, and Si. These elements are deposited in oxide or hydroxide oroxyhydroxide forms on a porous support such as silica, zirconia, andtitania. Supports can be pellets, granules, extrudates, spheres, and thelike in sizes of from about 1 to about 5 mm. The deposition can becarried out in a single step or multiple steps. The examples of themixed oxide catalysts are Nb₂O3-TiO₂, V2O₃—TiO₂, MoO₃—TiO₂, TiO₂—ZrO₂,Nb₂O₅—V2O₃, MoO₃—V₂O₅, MoO₃—ZrO₂, TiO₂—ZrO₂—SiO₂, TiO₂—Nb₂O₅—SiO₂,MoO₃—Nb₂O₅—TiO₂, V₂O₅—Nb₂O₅—TiO₂, MoO₃—Nb₂O₅—SiO₂, TiO₂—Bi₂O₃—SiO₂,MoO₃—NbO₅—ZrO₂, TiO₂—Nb₂O₅—Bi₂O3, MoO₃—V₂Os—TiO₂, TiO₂—Bi₂O₃—SiO₂,MoO₃—Bi₂O₃—SiO₂, TiO₂—ZrO₂—Bi₂O₃—SiO₂, and TiO₂—ZrO₂—Nb₂Os—Bi₂O₃—SiO₂.

The general procedure for preparing these mixed oxide catalysts areimpregnation and co-precipitation or a combination of these two, whichare performed in a single step or multiple steps. One may performimpregnation of one, two or three metal components on a porous supportor on a mixed oxide support prepared by co-precipitation. Impregnationcan be performed in one step or multiple steps.

Co-precipitation products and impregnation products obtained in powderyforms are subjected to suitable heat treatment at temperatures fromabout 150° to about 600° C. The powdery materials are shaped in asuitable size of from about 1 to 5 mm for the fixed bed reactor. Theshaped materials are calcined at temperature from 200° to about 750° C.,preferably from about 250° to about 600° C. in air. Optionally one ortwo metal components can be deposited on a shaped material prepared by aco-precipitation or impregnation method in either powder form or shapedform and then calcined at from 200° to about 750° C., preferably fromabout 250° to about 600° C. in air. The co-precipitation andimpregnation can be carried out in aqueous phase or in organic phase,such as hydrocarbons, ethers, ketones, alcohols, and mixtures of these.

When precipitation is carried out in organic phase, organometalliccompounds are preferably used. For example, two different solutions ofdifferent organometallic compounds are added to a suitable organicsolvent simultaneously with vigorous stirring under precipitationconditions at suitable temperatures. Sometimes a third solution isnecessary during the addition or afterward to cause gelation orprecipitation. An example of the third solution is water, basic oracidic water solution in a suitable organic solvent such as alcohol,ether, ketone, organic ester, or mixtures of these. Another optionalmethod is simultaneously adding first organometallic solution and thirdsolution to second organometallic solution with vigorous stirring. Ifnecessary, the precipitates are aged at a suitable temperature formabout 25° to about 200° C. for from 30 minutes to about 30 hours insuitable medium. Sometimes a co-precipitated product in aqueous mediumis aged in neutral, mildly acidic or basic organic medium.

The aging medium may or may not contain a minor amount of waterdepending on the nature of the material to be aged. The aging mediumcould be mildly acidic, mildly basic or neutral. The aged product isdried at a temperature from about 100° to about 400° C. and thencalcined at a temperature from about 250° to about 750° C. If necessary,impregnation of one or two elements on a suitable support is carried outby using an organic solution containing one or two organometalliccompounds or aqueous solution containing one or two compounds.Optionally one can perform multiple impregnations by using differentsolutions.

However, one may choose to use any heterogeneous catalyst disclosed inthe prior art, as long as the catalyst is suitable for a fixed bedreactor for the operation of a large commercial reactor. Examples ofheterogeneous catalysts disclosed in the prior art are titanium oxide,TS-1 , Ti-MCM-41 , molybdenum oxide, vanadium oxide, niobium oxide, leadoxide, and MgLa mixed oxide as appropriate and preferably supported asdescribed herein.

It is important that a support should have surface hydroxyl groups.Silica is a preferred support. The term “treated support” or “treatedsilica” is understood to mean a support having an optimized populationof surface hydroxyl groups for a given surface area for the preparationof the catalysts described herein. Depending on how silica is prepared,silica may not have a sufficient number of surface hydroxyl groups in agiven surface area. For such silica, the silica is treated to introduceextra surface hydroxyl groups with an aqueous base solution and thenwashed thoroughly with water, followed by calcination at a temperaturefrom 280° to 650° C., prior to use. Optionally one may attempt torehydrate a commercially available silica support. The rehydrated silicais calcined at a temperature from 280° to 650° C. to optimize thepopulation density of surface hydroxyl groups prior to use. Therefore,the preferred silica support used for the preparation of a solidcatalyst is “treated silica.” A class of preferred supports,particularly silica supports, is those that have had the surfacehydroxyl groups increased by treatment with a base solution as describeto obtain the maximum number of hydroxyl groups without degrading thephysical integrity and strength of the support. Controlling the sodiumcontent on silica support is very important for the preparation ofaromatic carbonates such as EPC and DPC, because basic impurities suchas alkali metal oxide on silica causes unwanted side reactions and tendsto cause catalyst instability. Alkali metal on silica support causesinstability of the catalyst performance and undesired side reactions forthe transesterification and disproportionation, which produces alkylaryl carbonate and diaryl carbonate. The preferred “treated silica”support will have less than about 0.05wt % Na, preferably less thanabout 0.03 wt % Na. Treating silica with aqueous alkali metal solutionhas an additional benefit of widening the pores. However, leaching outtoo much silica from silica support during the treatment with alkalimetal solution can cause the problem of maintaining physical integrityand strength.

Other catalysts disclosed in this invention are supported metal alkoxideor mixed metal alkoxide catalysts, which are prepared by bonding metalalkoxides to porous supporting materials through oxygen bridge bonds.The porous supporting materials must have surface hydroxyl groups, whichreact with alkoxy groups for the formation of oxygen bridge bonds. Thepreferred support is treated silica, which has less than about 0.05 wt %Na, preferably less than about 0.03 wt % Na. Optimizing the populationof the surface hydroxyl groups on silica support is very important tocreate stable, strongly anchored metal alkoxide active sites, since nohigh temperature calcinations is involved to link active metal atoms tothe surface of silica via M-O—Si bridge bonds. Maximizing the number ofM-O—Si bridge bonds is highly desirable.

Thus, an optimized population of hydroxyl groups for a given surfaceincludes the maximum number of hydroxy groups that can be obtained forthe given surface area within the constraints of alkali metal contentand support strength as described above.

The preferred metal alkoxides are Group IV and V metal alkoxides such astitanium alkoxides, zirconium alkoxides, vanadium alkoxides, niobiumalkoxides, and the like. The Group V metal alkoxides include a lowervalent alkoxide such as tetra-alkoxide and an oxytrialkoxide such asVO(OR)₃ or oligomers of oxoalkoxide. Examples of preferred supports aresilica, zirconia, titania, titania-silica, silica-alumina, andsilica-zirconia. Using treated silica is especially important for thepreparation of supported metal alkoxide catalysts. An alkoxide catalyston a support can have one or two different metal alkoxides.

The heterogeneous metal alkoxide catalysts are prepared by contacting ametal alkoxide solution or a mixed solution of two different metalalkoxides with a support such as silica at temperature from about 20° toabout 400° F., preferably from about 40° to about 300° F. The alkoxidesolution is prepared by dissolving one or two different metal alkoxidesin a solvent. The solvent must not interfere with the oxygen bridgeforming reaction in any way. Examples of such solvents are hydrocarbons,ethers, ketones, alcohols and mixtures thereof. When two different metalalkoxides are supported on a support, optionally two different metalalkoxide solutions are prepared and are reacted with a support insequence:

M(OR)_(n)+x OH(on a support surface)→(RO)_(n·x)(O—)_(x) (on thecatalyst)+x ROH

where n=4 or 5, x=1, 2, 3 or 4 and R=alkyl or aryl group

Supported titanium alkoxide is an acidic catalyst. The higher activityof supported titanium alkoxide catalyst on silica compared to titaniumalkoxide in homogeneous catalyst is attributed to higher acidity of thesupported Ti⁺⁴. The catalyst acidity plays an important role for acidcatalyzed transesterification and disproportionation reactions inproducing aromatic carbonates.

It is possible to prepare a supported metal alkoxide catalyst in situ asan optional method. Reactors are loaded with a treated support. Metalalkoxide solutions are circulated through the reactor at a temperaturefrom ambient to about 400° F. After formation of a supported metalalkoxide catalyst, any remaining solvent is drained off from reactors.After washing the reactors with a suitable solvent such as ethanol,pentane, or toluene and optionally heat treatment of the catalysts inflow of an inert gas such as nitrogen at a temperature from about 80° toabout 400° F., preferably from about 100° to about 350° F., thecatalysts are ready for the transesterification and disproportionation.For the preparation of supported metal or mixed metal alkoxidecatalysts, it is especially important to use a “treated support.” Anexample of a treated support is the treated silica described above.

In the present invention deactivation of heterogeneous catalysts wasobserved. This is especially true for the disproportionation reaction ofalkyl aryl carbonates such as MPC, EPC, etc. It was observed that thecause of catalyst deactivation is the deposition of heavy polymers onsolid catalyst, which blocks the catalyst active sites and fills up thecatalyst pores. The deactivated catalyst is either dark brown or black,depending on the degree of polymer deposition. The deactivation rate fora given heterogeneous catalyst is much faster for a disproportionationreaction than a transesterification reaction. It was discovered thatdeactivated catalyst could be regenerated in situ by depolymerizing thepolymers deposited on the catalyst by contacting the deactivatedcatalyst with a compound containing a hydroxyl group, such as steam,methanol, ethanol, phenol, or mixtures of hydroxy compounds at anelevated temperature. Optionally one can use a solution of hydroxycompound in a solvent. The preferred solvents are benzene, toluene,xylenes, pentane, hexane, octane, decane, THF or any mixtures ofsolvents. The use of a solvent for the depolymerization of polymersdeposited on deactivated catalyst is only optional and not required. Theregenerated catalyst is preferably dried at temperature from 200° toabout 500° F. in an inert gas (e.g. nitrogen) flow prior to use. Thecatalyst regeneration with hydroxy containing compounds was performed bytreating deactivated catalysts in situ in the flow of an alcohol orwater-alcohol solution or with steam at a temperature from 250° to 600°F., preferentially from 270° to 450° F. The treatment of the deactivatedcatalysts may be performed with both alcoholic solution and steam. Forexample, a deactivated catalyst may be treated first with alcohol oralcohol solution and then with steam or in reverse order. It ispreferable to perform the catalyst regeneration under sufficientpressure so that at least, some liquid phase present in the catalystbed. The preferred alcohol is methanol, ethanol or mixture of these two.Optionally an alcohol solution may contain water in amount of up to 80wt %, preferably up to 20 wt %, most preferably up to 5%. The streamfrom the depolymerization reactor contains phenol, DEC, EPC and traceamounts of phenetole and heavies as depolymerization reaction products.

To use a glycol ether to wash off the polymers on the catalyst andimprove selectivity, the solvent as described above, is introduced as acomponent of reaction mixture to a catalytic reaction zone and separatedfrom reaction mixture for recycling. The examples of such ether solventsare ethylene glycol diethyl ether, ethylene glycol dimethyl ether,diethylene glycol diethyl ether, diethylene glycol dimethyl ether, etc.

The catalyst regeneration technique disclosed in this invention can alsobe used in any process for the production of aromatic carbonates, wherea homogeneous catalyst is used. For the regeneration of a homogeneouscatalyst system, an alcohol solution must be fairly dry so that watercontent may not exceed about 0.2 wt %. Therefore, the catalystregeneration technique disclosed in this invention is applicable for anyprocess for producing organic carbonates.

It was surprisingly discovered that this catalyst deactivation could bealleviated by performing the disproportionation reaction in the presenceof an aromatic hydroxy compound such as phenol. Furthermore, there is anadditional benefit by the performing disproportionation in the presenceof a trace amount of water in the catalytic reaction zone in an amountof up to 0.3 wt %, preferably up to 0.10%. It appears that a substantialportion of heavy polymers deposited on the catalysts are polycarbonates.Performing disproportionation in the presence of an aromatic hydroxylcompound or both aromatic hydroxyl compound and a trace amount of waterin the catalytic reaction zone results in acceptable stable catalystperformance. However, it is understood that too much of an aromatichydroxy compound results in unacceptably low rate of producing DPC dueto the equilibrium nature of the transesterification reaction.Therefore, maintaining the mole ratio of aromatic hydroxy compound toDPC from 0.05 to 10, preferably from 0.1 to 6, in the disproportionationreaction zone is essential to ensure both long acceptable catalyst cycletime and good productivity of DPC. For the transesterification reaction,the mole ratio of phenol to DEC is maintained higher than 0.2,preferably higher than about 0.3, most preferably higher than about0.35.

An example of the schematic process flow diagram for the production ofDPC is illustrated in the FIGURE. There are two fixed bed, dual phasereactors and five distillation columns. The first dual phase reactor 38is primarily for the transesterification reaction to produce EPC fromDEC and phenol. The second dual phase reactor 37 is primarily for thedisproportionation reaction to produce DPC from EPC. Both dual phasereactors 38 and 37 are loaded with either a heterogeneous catalyst oroptionally two different heterogeneous catalysts disclosed in thisinvention. Optionally one may introduce additional fixed bed reactors(not shown) in series between 38 and 37. The major objective of thisadditional reactor is to produce additional EPC. The fresh phenol feedstream 1 and fresh DEC feed stream 2 are mixed with recycle streams 11and 25, and then introduced to the dual phase reactor 38 through line 3.Optionally a nitrogen gas stream is introduced to 38 via line 4. Theproduct EPC and the co-product ethanol are produced by performing thetransesterification in 38. In the reactor 38, the co-product ethanol isvaporized into gas phase. The reactor effluent stream 5 is introduced tothe first distillation column 30 via lines 5, 6 and 7. Ethanol isstripped off in the column 30 to overhead stream 8 with small amount ofDEC. Stream 8 is partially cooled and then introduced to gas-liquidseparation drum 36, and the gas stream 9 is recycled to the Reactor 38and 37. The liquid stream 10 (ethanol) from 36 is recycled to DEC plantto produce DEC. A side-draw stream 11 from column 30, which is composedof DEC, phenol and EPC is recycled back to the reactor 38 through line3. Optionally one may choose to place a catalyst bed 101 at the bottomsection of the column 30 below the sidedraw point for the additionalconversion of phenol to EPC. The distillation column 30 is designed andoperated so that the side-draw stream 11 is essentially free of ethanol.The recycle loop for the dual phase reactor 38 comprises the lines 5, 6and 7, column 30, and lines 11 and 3. The bottom stream 12 from thecolumn 30 is introduced to the second dual phase reactor 37 via lines 13and 14. The stream 12 contains EPC and phenol. But it also containssmall amounts of DPC and the byproduct phenetole. The column 30 is alsodesigned and operated to minimize the DEC in the bottom stream 12. Thestream 12 is combined with recycle stream 19 to stream 13. The stream 13is combined with nitrogen gas stream 15 to the stream 14, which isintroduced to 37. The product DPC and the co-product DEC are produced inthe reactor 37 by performing disproportionation of EPC. DEC in thereactor 37 is vaporized to gas phase. The reactor effluent steam 16 isintroduced to the second distillation column 32. The stream 16 iscomposed of mostly DEC, EPC, phenol, DPC, and small amounts of ethanoland by-products. Ethanol and DEC in stream 16 are stripped off alongwith vapor in the column 32 as overhead stream 17, which is introducedto the first column 30 via line 7. The bottom stream 18 from the column32 splits to two streams 19 and 20. The stream 19 is recycled back tothe second reactor 37 via lines 13 and 14. The recycle loop for thesecond dual phase reactor 37 comprises the line 16, column 32, and lines18, 19, 13 and 14. The other stream 20 is introduced to the thirddistillation column 34, where remaining DEC in the stream is recoveredas overhead stream 21, which is sent to the first column 30 via thelines 6 and 7. A side-draw stream 22 from 34 is introduced to the fourthdistillation column 35 to remove the byproduct phenetole as the overheadstream 24. The bottom stream 25 from the column 35 is recycled to 38.Optionally the bottom stream 25 from 35 may be recycled to the firstcolumn 30 via lines 31, 21, 5, 6, and 7. The bottom stream 23 from 34 isintroduced to distillation column 33 to recover the product DPC. Theoverhead stream 26, which is mainly composed of EPC, is recycled to thedisproportionation reactor 37 via line 14. The column 33 is operatedunder sub-atmospheric pressure. The bottom stream 27 from 33 is thecrude DPC stream. One may choose to use other material, such diethylether, dimethyl ether, isopentane or butane, in place of nitrogen gas orto partially replace the nitrogen gas for the operation of the two dualphase reactors, 38 and 37.

Control Example 1

A titanium oxide (9.2 wt %) catalyst supported on silica was preparedaccording to the prior art (WO 03/066569). 3.83 g Ti(OC₄ H_(9·n))₄ wasdissolved in 90 ml dry toluene. Granular silica (+8 mesh, 655 ppm Na byweight, 300 m2/g BET SA and 1 cc/g PV) was pre-dried at 330° C. for 2hrs in air. The titanium butoxide solution was refluxed with 25 ml (9.15g) of the dried silica granules in boiling toluene solution in a 200 mlflask with a condenser. After about 6 hours reflux, the excess toluenein the flask was boiled off from the flask. The titanium butoxidesupported silica was recovered from the flask and dried at 120° C. in avacuum oven for 1.5 hours. The dried silica was calcined at 500° C. forto 2 hours. The calcined product was paper white granule. The totalcatalyst weight was 9.67 g. This is catalyst A.

Example 2

A mixed niobium/titanium oxide catalyst supported on silica was preparedaccording to the invention. 0.592 g of Nb(OC₄ H_(9·n))₅ was dissolved in80 ml toluene. The granular silica used in the Control Example 1 isdried at 320° C. for 2 hours in air. 25 ml (9.27 g) of this dried silicais refluxed with the above niobium butoxide solution. After 5.5 hoursreflux in the same fashion as the Control Example 1, the excess toluenewas drained out from the flask. A water solution in methanol prepared bymixing 0.209 g of water with 120 ml methanol was poured into flask andthen the solution was refluxed in boiling methanol for an hour. Theexcess of methanol was drained out from the flask. A titaniumtetrabutoxide solution prepared by dissolving 3.56 g of titaniumbutoxide in 90 ml toluene was poured into the flask and then the contentin the flask was refluxed in boiling toluene for 5.5 hours. The excesstoluene in the flask was boiled off from the flask. The material in theflask was recovered from the flask and dried at 120° C. in a vacuum ovenfor 1.5 hours. The dried silica was calcined at 500° C. for to 2 hours.The appearance of the calcined product was different from the catalystin the Control Example 1. It looked more like granular silica supportthan the paper white catalyst in the Control Example 1. The totalcatalyst weight was 10.28 g. This is catalyst B.

Example 3

In this example, treated silica was used to prepare a mixedtitanium/niobium oxide catalyst supported on silica. A mixedniobium/titanium oxide catalyst supported on treated silica was preparedaccording to the invention. The same granular silica of Control Example1 was used to prepare the treated silica of this example. The granularsilica (40.56 g) was treated with a sodium hydroxide solution preparedby dissolving 8.05 g NaOH in 226 g water at room temperature for 7minutes with stirring. The treated silica was washed with cold waterthoroughly and then with hot water (about 65° C.) several times toremove trace of sodium on silica. The treated silica was dried at 150°C. for 2 hours and then calcined at 325° C. for 2 hours. This calcinedsilica contained 300 ppm Na by weight. A niobium alkoxide solution wasprepared by dissolving 0.844 g of Nb(OC₄H_(9·n))₅ in 80 ml toluene. 8.46g of the treated silica was refluxed in the above niobium butoxidesolution for 3 hours in a flask with water-cooled condenser. Aftercooling, the excess solution in the flask was drained out of the flask.A water-methanol mixture was prepared by mixing 0.645 g water with 90 mlmethanol. This water-methanol mixture was poured into the flask and thecontent in the flask was again refluxed. After an hour reflux, theexcess solution in the flask was drained out. A titanium tetrabutoxidesolution prepared by dissolving 3.67 g of titanium butoxide in 80 mltoluene was poured into the flask and then the content in the flask wasrefluxed for 1 hour 45 minutes. The excess toluene in the flask wasdistilled off from the flask. The material in the flask was recoveredfrom the flask and dried at 120° C. in a vacuum oven for 1 hour. Thedried silica was calcined at 500° C. for to 2 hours. The appearance ofthe catalyst was more like granular silica support. This is catalyst C.

Example 4

A titanium alkoxide catalyst supported on treated silica was preparedaccording to the invention. The same granular silica (80.55 g) used inthe Control Example 1 was treated with a sodium hydroxide solutionprepared by dissolving 8.3 g NaOH in 580 ml of deionized water at roomtemperature for 8 minutes with stirring. The treated silica was washedwith cold water and then hot water. The washed silica was treated withammonium nitrate solution prepared by dissolving 99 g ammonium nitratein 2 liter deionized water at 80° C. for 2 hours. The treatment ofsilica with ammonium nitrate solution was repeated 13 times. Finally thesilica washed with deionized water at ambient temperature. The washedsilica was dried at 110° C. for an hour, followed by calcination at 370°C. for 1.5 hours and then 375° C. for 30 min. The calcined silicacontained 23 ppm Na by weight and 2.9% weight loss on calcinations at550° C. A titanium tetrabutoxide was prepared by dissolving 4.74 gtitanium tetrabutoxide in 70 ml toluene. 10.44 g (31 ml) of treatedsilica was refluxed in the titanium tetrabutoxide solution with for 6hours and then the excess titanium solution in the flask was drainedout. The drained titanium solution contained 0.50 wt % Ti. The silicawas washed with 90 ml toluene at room temperature. The washed productwas dried at 170° C. for 3 hours in a vacuum oven. The appearance of thefinished catalyst looked like treated silica granules. A small portionof this finished catalyst was calcined at 500° C. for 2 hrs in air todetermine Ti content on the catalyst. The Ti content of the calcinedcatalyst is 3.25%. The appearance of the calcined catalyst looked liketreated silica granules. This experiment indicates a successful supportof titanium alkoxide on treated silica.

Performing Transesterification and Disproportionation

The catalysts were tested in a unit having a fixed bed reactor, adistillation column and a reflux drum. The dimension of the fixed bedreactor was ½ inches diameter and 15 inches long. It had threethermocouples to monitor temperatures at three positions just above thecatalyst bed, the middle of catalyst bed, and just below the catalystbed. The top half and bottom half of the reactor temperatures areindependently controlled. The unit also had a feed preheater at the topof the fixed catalyst bed, whose temperature is separately controlled.The fixed bed reactor was operated in down flow mode. The distillationcolumn consists of a 2 liters capacity reboiler and 42 inch by 1 inch OD(0.870 inch ID) column. There were three pressure transmitters tocontrol and record the column overhead pressure, the top and bottom ofthe catalyst bed. Also the reboiler had a liquid level transmitter. Thefixed bed had a recirculation loop; the stream to the reactor from theliquid medium in the reboiler was pumped through the reactor in downflow and then returned to the reboiler. The fresh feed DEC was pumpedinto the recirculation loop prior to the reactor. The fresh phenol feedsolution was separately pumped into the recirculation loop prior to thereactor. Nitrogen gas was introduced into the system as necessary. Thevapor from the distillation column was condensed and removed as anoverhead liquid stream. Non-condensable gas was discharged from thecondenser.

For the continuous run for the transesterification, the phenol feedsolution was continuously fed at a given rate, while continuouslyremoving product stream from the reboiler at a predetermined constantrate and continuously removing overhead product. The DEC feed rate is acascade to maintain a constant liquid level in the reboiler.

Test 1

The purpose of this run is to demonstrate the performance of a dualphase fixed bed reactor for the transesterification reaction. The dualphase in the catalytic reaction zone was created by a boiling mixture ofvapor phase and liquid phase in the catalytic reaction zone.

The catalyst A (25 ml; 9.67 g) was loaded in the reactor. 167.6 g phenoland 737.4 g DEC were charged in the reboiler. The testing was performedat the conditions listed in Table 1:

TABLE 1 Overhead Column Pressure, psig 18 Reboiler temperature, ° F. 335Distillation column temperature, ° F. 300-310 Recirculation rate, ml/min66 Pressure at top of fixed bed reactor, psig 24.6 Pressure at bottom offixed bed reactor, psig 20.5 Fixed bed reactor temperature, ° F. 338-342Nitrogen flow rate to the reboiler, ml/min 60 Reflux from reflux drum,ml/min 0

Run hours started when the fixed bed reactor temperature reached thetarget temperature of 340° F. During the run, DEC was continuouslypumped in to maintain a constant liquid level. After running for 45hours, 23 g phenol was charged into the system as a 20.92 wt % phenolsolution over a period of 55 minutes at 2 ml/min flow rate. The run wascontinued to 77 hours on stream time. The average overhead liquid flowrate was about 0.3 ml/min. The overhead stream was composed mainly ofDEC and a small amount of ethanol. Samples were taken from the reboilerand overhead stream for analysis. At the end of 77 hours run time, 64.8mole % of phenol charged to the unit had been converted. The yields were63.1 mole % EPC; 1.52 mole % DPC and 0.26 mole % by-products based onthe total amount of phenol charged into the system. Phenetole was themajor by-product, which accounted for 33.6 mole % of byproducts. Theaverage productivity was 1.81 m/h/kg of catalyst for EPC; 0.041 m/h/kgof catalyst for DPC and 0.01 m/h/kg of catalyst for by-products. Theproductivity of EPC in this example is superior to those disclosed inprior art, demonstrating the superior productivity of the processdisclosed in this invention.

Test 2

The purpose of this run is to demonstrate the performance of a dualphase fixed bed reactor and a mixed niobium and titanium oxide catalystaccording to the invention for the transesterification reaction. Thedual phase in the catalytic reaction zone was created by a boilingmixture of vapor phase and liquid phase in the catalytic reaction zone.

Catalyst B (24 ml; 9.172 g) was loaded in the reactor. This test wasperformed under the identical condition to Test 1. 167.6 g phenol and737.4 g DEC were charged in the reboiler. Run hours started when thefixed bed reactor temperature reached the target temperature of 340° F.During the run, DEC was continuously pumped in to maintain a constantliquid level. After running for 22 hours, 91.62 g phenol was chargedinto the system as 31.38 wt % phenol solution over a period of 2 hours20 minutes at 2ml/min flow rate. The run was continued to 73 hours onstream time. The average overhead liquid flow rate was about 0.3 ml/min.The overhead stream was composed mainly of DEC and a small amount ofethanol. Samples were taken from the reboiler and overhead stream foranalysis. At the end of 73 hours run, 64.6 mole % of phenol charged tothe unit had been converted. The yields were 63.1 mole % EPC; 1.4 mole %DPC; and 0.27 mole % by-products based on the total amount of phenolcharged into the system. Phenetole was the major byproduct, which wasaccounted for 55.9 mole % of by-products. The average productivity was2.567 m/h/kg of catalyst for EPC; 0.059 m/h/kg of catalyst for DPC and0.022 m/h/kg of catalyst for by-products. The major by-product wasphenetole, which was accounted for 33.6 mole % of the total by-products.The productivity of EPC in this example is superior to the catalystdisclosed in prior art and Test 1.

Test 3

This test was performed to demonstrate the disproportionation of EPC toDPC in batch operation. The fixed bed reactor was operated in boilingpoint mode by using the same equipment as described.

A crude feed was prepared as in the Test 2 and then the excess of DEC inthe crude feed was distilled off to concentrate EPC and then toluene wasadded to it. The purpose of toluene addition to the concentrated crudefeed was to perform the disproportionation under the dual phase mode inthe down flow reactor. Boiling reaction mixture of vapor phase andliquid phase were created in the reaction zone. Run hours started whenthe fixed bed reactor temperature reached the target temperature of 340°F. The test was performed with the same catalyst used in the Test 2.During the run, toluene was continuously pumped in to maintain aconstant liquid level and the overhead flow rate was about 7 ml/min. Theoverhead stream was composed mainly of toluene, a small amount of DECand a trace amount of ethanol. The total weight of the feed in thereboiler was 635.3 g. The composition of the feed was 44.19 wt %toluene; 8.42 wt % DEC; 0.07 wt % phenetole; 0.29 wt % by-products;11.21 wt % phenol; 31.88 wt % EPC and 3.94 wt % DPC. The testing wasperformed at the conditions as given in Table 2:

TABLE 2 Overhead Column Pressure, psig 16.3 Reboiler temperature, ° F.311 Distillation column temperature, ° F. 295-305 Recirculation rate,ml/min 67 Pressure at top of fixed bed reactor, psig 32.8 Pressure atbottom of fixed bed reactor, psig 18.5 Fixed bed reactor temperature, °F. 327-330 Nitrogen flow rate to the reboiler, ml/min 50 Reflux fromreflux drum, ml/min 0

After 6 hours operation, the analysis of product from reboiler indicated5.8 mole % conversion of EPC with 95.2 mole % selectivity to DPC. Theproductivity of DPC was 0.614 m/h/kg of catalyst.

Test 4

This test was performed to demonstrate a continuous run for theproduction of EPC with a fixed bed reactor operated in boiling pointreactor mode under a steady state condition.

Another batch of the catalyst identical to the catalyst B in the Example2 was prepared. 9.04 g (about 25 ml) of this catalyst was loaded in thereactor. 200 g phenol and 610 g DEC were charged in the reboiler for acontinuous run. Run hours started when the fixed bed reactor temperaturereached the target temperature of 340° F. During the run, DEC wascontinuously pumped in to maintain a constant liquid level. Afterrunning for 23.25 hours, 85.58 g phenol was charged into the system as31.38 wt % phenol solution over a period of 2 hours 15 minutes at 2ml/min flow rate. The continuous run at a steady state condition wasperformed by continuously pumping in 32.2 wt % phenol solution in DEC at0.18 ml/min and continuously removing the product stream at 0.15 ml/min.DEC was pumped in to maintain a constant liquid level 81% in reboiler.At 242.5 hrs on stream time under at steady state operation condition,the following results, given in Tables 3 and 4, was obtained:

TABLE 3 On stream time, hrs 242.5 Overhead Column Pressure, psig 19.6Reboiler Temperature, ° F. 354 Distillation column temperature, ° F. 330Recirculation rate, ml/min 67 Pressure at top of fixed bed reactor, psig20.8 Pressure at bottom of fixed bed reactor, psig 19.2 Fixed bedreactor temperature, ° F. 341 Nitrogen flow rate to the reboiler, ml/min60 Reflux from reflux drum, ml/min 0 Liquid level in reboiler, % 81.04Feed rate of 32.2 wt % PhOH/DEC solution, ml/min 0.15 DEC flow rate,ml/min 0.28 Overhead Flow Rate, ml/min 0.269 Bottom product flow rate,ml/min 0.185

TABLE 4 Composition of Streams (wt %): Component Overheads BottomsEthanol 3.5793 0.0448 DEC 96.1344 62.5443 Phenetole — 0.0176 Phenol0.2208 17.5226 EPC 0.0655 19.2401 DPC — 0.6306 Phenol Conversion (mole%) 39.5 EPC Yield (mole %) 37.6 EPC Selectivity (mole %) 95.04 DPC Yield(mole %) 1.9 DPC Selectivity (mole %) 4.87 Phenetole Selectivity (mole%) 0.05 EPC Productivity (mol/h/kg catalyst) 1.5

Phenetole was the only by-product detected in the product stream by GCand GC-mass spectrometry (GC-MS). This is a result superior to priorart. Deactivation of the catalyst was observed.

Test 5

This test was performed to demonstrate a continuous run for thetransesterification reaction between propylene carbonate and ethanol toproduce DEC and propylene glycol co-products by using a fixed bedreactor operated in boiling point mode.

Another batch of the catalyst identical to the catalyst B in the example2 was prepared. 9.6 g (about 25 ml) of this catalyst was loaded in thereactor. 280 g propylene carbonate and 635 g ethanol were charged in thereboiler for a continuous run. Run hours were started when the fixed bedreactor temperature reached the target temperature of 335° F. During therun, the product stream was continuously removed from the reboiler at aconstant rate of 0.14 ml/min. To maintain a constant liquid level of85%, a 23.1 wt % propylene carbonate solution in ethanol wascontinuously pumped in to the reactor cascade mode. The run wascontinued for 162 hours. During the run, the conversion of propylenecarbonate was fairly constant. The average conversion of propyleneduring the run was 14.5 mole %. The average DEC productivity was 1.81mole/h/kg of catalyst.

Test 6

The objective of this experiment is to demonstrate a superiorperformance of the mixed oxide catalyst supported on treated silica tothe conventional catalyst disclosed in prior art. A slower deactivationrate, higher activity of the Catalyst C and the catalyst regeneration bydepolymerization of deactivated Catalyst C were demonstrated. The testof a conventional titanium oxide catalyst supported on silica (TheControl Example 1) was performed in Test 6A. The test of the mixed Ti/Nboxide catalyst supported on treated silica (Catalyst C in the Example 3)was performed in the Test 6B.

Test 6A

Another Ti oxide catalyst similar to the Catalyst A in the ControlExample 1 was prepared. The only difference from the Control Example 1was the calcination of the same silica at 340° C. for 4 hours in air. 25ml (9.0 g) catalyst was loaded in the fixed bed reactor described. Theappearance of this catalyst is white sugar coated silica granules thesame as the catalyst of the Control Example 1.285 g phenol and 530 g DECwere charged in the reboiler. The testing was performed at the followingconditions listed in Table 5:

TABLE 5 Overhead Column Pressure, psig 19-21 Reboiler Temperature, ° F.345-347 Distillation column temperature, ° F. 270-285 Recirculationrate, ml/min 66-67 Pressure at top of fixed bed reactor, psig 20-22Pressure at bottom of fixed bed reactor, psig 19-22 Fixed bed reactortemperature, ° F. 332-337 Nitrogen flow rate to the reboiler, ml/min 60Reflux from reflux drum, ml/min  0 Feed rate of 32.2 wt % PhOH/DECsolution, ml/min 0.17-0.18 Bottom product flow rate, ml/min 0.18-0.19

During the run, the flow of DEC was cascaded to a liquid level of 88%(arbitrary scale) in to maintain a constant liquid level. The run wasperformed by continuously draining the reaction mixture in the reboiler,while pumping in 40.4 wt. % phenol solution in DEC at a constant rate ofabout 0.19 ml/min. The run was continued for 115 hours withoutinterruption. The overhead stream was composed of mostly DEC, ethanoland small amounts of phenol and EPC. Samples were taken from thereboiler and overhead stream for analysis. Phenetole was the onlybyproduct in the bottom product stream from the reboiler. The resultsare listed in Table 6.

TABLE 6 Hours on stream time 31 61 85 115 Phenol Conversion (mole %)17.8 20.8 17.7 11.9 EPC Productivity 0.771 0.869 0.754 0.502 (mol/h/kgcatalyst) DPC Productivity 0.008 0.013 0.013 0.0092 (mol/h/kg catalyst)Phenetole productivity 0.0016 0.0018 0.0018 0.0017 (mol/h/kg catalyst)

Test 6B

The performance of the Catalyst C prepared in the Example 3 was testedfor the transesterification of DEC with phenol. 8.4 g (25 ml) catalyst Cwas loaded in the fixed bed reactor. 287 g phenol and 530 g DEC werecharged in the reboiler. The testing was performed at the conditionslisted in Table 7:

TABLE 7 Overhead Column Pressure, psig 17-19 Reboiler Temperature, ° F.345-348 Distillation column temperature, ° F. 270-285 Recirculationrate, ml/min 66-67 Pressure at top of fixed bed reactor, psig 19-21Pressure at bottom of fixed bed reactor, psig 19-22 Fixed bed reactortemperature, ° F. 331-335 Nitrogen flow rate to the reboiler, ml/min 60Reflux from reflux drum, ml/min  0 Feed rate of 32.2 wt % PhOH/DECsolution, ml/min 0.18-0.19 Bottom product flow rate, ml/min 0.18-0.20

During the run, the flow of DEC was cascaded to a liquid level of 88%(arbitrary scale) in to maintain a constant liquid level. The run wasperformed by continuously draining the reaction mixture in the reboiler,while pumping in 36.5 wt % phenol solution in DEC. The result of the 234hours time-on stream is listed in Table 8. The overhead stream wascomposed of mostly DEC, ethanol and small amounts of phenol and EPC.Samples were taken from the reboiler and overhead stream for analysis.Phenetole was the only by-product in the bottoms stream from reboiler.The result of this run in Table 8 clearly indicates slower deactivationand higher activity than the Catalyst A of the Test 6A.

The result of this run, given in Table 8, clearly indicates slowerdeactivation and higher activity of the Catalyst C than the Catalyst A.

TABLE 8 Hours on stream time 33 56 94 160 216 234 Phenol Conversion(mole %) 23.5 38.5 34.7 31.9 29.1 25.0 EPC Productivity (mol/h/kgcatalyst) 0.686 1.119 1.027 0.976 0.556 0.510 DPC Productivity (mol/h/kgcatalyst) 0.008 0.020 0.019 0.019 0.012 0.012 Phenetole productivity(mol/h/kg catalyst) 0.0018 0.0019 0.0014 0.0011 0.0006 0.0006

The run was continued to 360 hours of on-stream time. Continued catalystdeactivation was observed. At an on stream time of 360 hours, thereactor operation for the synthesis of aromatic carbonate was terminatedto perform catalyst regeneration. The catalyst regeneration wasperformed at 340° F. and 230 psig by depolymerizing polymers on thecatalyst by circulating ethanol through the reactor for 17 hours. Themajor products of depolymerization reaction with ethanol were phenol,DEC and unidentified by products. The run was resumed with theregenerated catalyst. The test result of the regenerated catalyst islisted in Table 9.

The result in Table 9 clearly demonstrates that the catalyst can beregenerated by performing depolymerization of the polymers deposited onthe catalyst.

TABLE 9 Hours on stream time 360 (before 393 (after regeneration)regeneration) 428 Overhead Column pressure, psig 18.2 19.5 18.1 ReboilerTemperature, ° F. 347 344 344 Distillation column temperature, 273-285273-286 271-285 ° F. Recirculation rate, ml/min 65 66 67 Pressure at topof fixed bed 20 22.2 20.4 reactor, psig Pressure at bottom of fixed bed19.2 22.1 20 reactor, psig Fixed bed reactor temperature, ° F. 330330-331 330-331 Feed rate of 32.2 wt % PhOH/DEC 0.18 0.18 0.18 solution,ml/min Bottom product flow rate, ml/min 0.16 0.18 0.18 Phenol Conversion(mole %) 17.5 23.4 27.2 EPC Yield (mole %) 16.47 22.8 26.2 EPCSelectivity (mole %) 94.35 97.3 96.3 DPC Yield (mole %) 0.95 0.610 1.017DPC Selectivity (mole %) 5.44 2.60 3.73 Phenetole Selectivity (mole %)0.21 0.13 0.10 EPC Productivity 0.452 0.681 0.830 (mol/h/kg catalyst)Phenetole productivity 0.013 0.009 0.016 (mol/h/kg catalyst)

While the disclosure includes a limited number of embodiments, thoseskilled in the art, having benefit of this disclosure, will appreciatethat other embodiments may be devised which do not depart from the scopeof the present disclosure. Accordingly, the scope should be limited onlyby the attached claims.

1. A solid disproportionation or transesterification catalystcomposition selected from the group consisting of oxides, hydroxides,oxyhydroxides and alkoxides of two to four elements from Group IV, V,and VI of the Periodic Table supported on a porous material which hassurface hydroxyl groups.
 2. The solid catalyst composition according toclaim 1 wherein said porous material has been subjected to a treatmentto increase the number of hydroxyl groups thereon.
 3. The solid catalystcomposition according to claim 2 wherein said treatment comprisescontacting said porous material with a basic solution.
 4. The solidcatalyst composition according to claim 3 wherein said porous materialcomprises silica.
 5. The solid catalyst composition according to claim1, wherein said alkoxide comprises a metal alkoxide.
 6. The solidcatalyst composition according to claim 1, wherein said alkoxidecomprises a mixed metal alkoxide.
 7. The solid catalyst compositionaccording to claim 1, wherein said alkoxide comprises an alkoxide ormixed metal alkoxide of Group IV and V metals.
 8. The solid catalystcomposition according to claim 1, wherein in the solid catalyst isTiO₂/Nb₂O₅ supported on treated silica, said treated silica having lessthan 0.05 wt. % Na.
 9. The solid catalyst composition according to claim1, wherein said catalyst composition comprises two differentheterogeneous catalyst compositions selected from the group consistingof oxides, hydroxides, oxyhydroxides and alkoxides of two to fourelements from Group IV, V, and VI of the Periodic Table supported onporous material which have surface hydroxyl groups.